Process for preparing 1,4-butanediol by catalytic hydrogenation of 1,4-butinediol

ABSTRACT

A process for preparing 1,4-butanediol by continuous catalytic hydrogenation of 1,4-butynediol comprises reacting 1,4-butynediol with hydrogen in the liquid continuous phase in the presence of a heterogeneous hydrogenation catalyst at from 20 to 300° C., a pressure of from 1 to 200 bar and values of the liquid-side volumetric mass transfer coefficient k L a of from 0.1 s −1  to 1 s −1    
     a) using a catalyst suspended in the reaction medium, where if a packed bubble column is employed this is operated in the upflow mode and at a ratio of gas leaving the reaction vessel to gas fed to the reaction vessel of from 0.99:1 to 0.4:1, or 
     b) passing the liquid and gas in cocurrent in an upward direction through a fixed-bed reactor operated in the gas-circulation mode while maintaining a ratio of the gas fed to the reaction vessel to gas leaving the reaction vessel of from 0.99:1 to 0.4:1.

The present invention relates to a process for preparing 1,4-butanediolby catalytic hydrogenation of 1,4-butynediol with hydrogen in thepresence of a solid hydrogenation catalyst at a pressure of from 1 to200 bar and values of the volumetric liquid-side mass transfercoefficient k_(L)a of from 0.1 s⁻¹ to 1 s⁻¹, where the liquid forms thecontinuous phase and the hydrogen forms the dispersed phase.

The hydrogenation of 1,4-butynediol to give 1,4-butanediol via theindividual steps shown in simplified form in the following scheme

has been carried out for decades and has been described many times.However, the known processes have the disadvantages of a low,uneconomical space-time yield (STY), ie. the amount of starting materialused per reactor volume and unit time, when hydrogenation is carried outat pressures below 200 bar, low catalyst lives and low selectivity. Inaddition, when fixed-bed catalysts are used, the hydrogenation requiresa high pressure of over 200 bar which is associated with high capitalcosts.

Furthermore, 1,4-butynediol, 1,4-butenediol and compounds derivedtherefrom, eg. the acetal from butanediol and hydroxybutyraldehyde whichis formed by isomerization of butenediol, can be separated bydistillation from 1,4-butanediol only with difficulty, if at all.However, for the further processing of 1,4-butanediol, it is criticalfor most applications that no incompletely hydrogenated products arepresent therein.

In chemical reactions, the selectivity generally decreases withincreasing conversions. Efforts are therefore made to carry out thereaction, on the one hand, at as low as possible a temperature and, onthe other hand, with a partial conversion in order to obtainselectivities which are as high as possible. In the hydrogenation ofbutynediol, complete conversion is essential with regard to the productquality achievable on work-up and the hydrogenation is therefore oftendistributed over a plurality of reactors which are operated underdifferent conditions.

U.S. Pat. No. 5,068,468 discloses the hydrogenation of 1,4-butynediolover solid supported nickel/copper catalysts in which space-time yieldsof 0.3 kg of butanediol/l·h at a pressure of 250 bar.

BE-B 745 225 describes the use of fixed-bed Raney nickel catalysts at259 bar, which achieve a space-time yield of 0.286 kg of butanediol/l·hin a two-stage process.

U.S. Pat. No. 4,153,578 discloses a two-stage process for thehydrogenation of 1,4-butynediol over suspended Raney nickel/molybdenumcatalysts at a pressure of 21 bar. This process achieves space-timeyields of 0.06 kg of butanediol/l·h.

DD-A 272 644 describes the suspension hydrogenation of aqueousbutynediol over nickel/SiO₂ catalysts. Assuming that butynediol is asusual used in a concentration of from 39 to 50% by weight and assumingcomplete conversion, the space-time yield is calculated as from 0.15 to0.25 kg of butanediol/l·h at a pressure of 15 bar. The catalyst useddisplays a loss in activity of 37% after only 50 hours.

For Example 1 of U.S. Pat. No. 2,967,893, a space-time yield of about0.01 kg of butanediol/l·h can be calculated for the Raneynickel-copper-catalyzed hydrogenation of 1,4-butynediol.

RU-A 202 913 describes the hydrogenation of butynediol over anickel/chromium catalyst at a space-time yield of 0.1 kg ofbutanediol/l·h.

EP-B 0 319 208, DE-A 19 41 633 and DE-A 20 40 501 disclose, inter alia,general hydrogenation processes which can be applied to 1,4-butynedioland in which the gas-circulation operating mode of the reactor isavoided by gas and liquid phases flowing in cocurrent from the bottomupwards through a fixed-bed catalyst. Here, gas and liquid phases flowthrough the catalyst in the form of the transition stream, with theliquid phase forming the continuous phase.

However, these processes have the disadvantage that in the case of highbutynediol loadings in the hydrogenation feed the reaction mixture atthe end of the reaction zone is depleted in hydrogen and as a resultonly an incomplete conversion of the 1,4-butynediol is achieved, thusleading to intermediates which can be separated from butanediol onlywith difficulty, if at all.

In the case of lower butynediol loadings, a complete conversion andsatisfactory product quality can be achieved only if a significantlyreduced space-time yield or higher operating pressure is accepted.

It is an object of the present invention to provide a process for thecatalytic hydrogenation of 1,4-butynediol to 1,4-butanediol by means ofwhich a high space-time yield together with high selectivity and highcatalyst operating lives can be achieved at a pressure of below 200 bareven when using technical-grade 1, 4-butynediol.

We have found that this object is achieved by a process for preparing1,4-butanediol by continuous catalytic hydrogenation of 1,4-butynediol,which comprises reacting 1,4-butynediol with hydrogen in the liquidcontinuous phase in the presence of a hydrogenation catalyst at from 20to 300° C., preferably from 60 to 220° C. and particularly preferablyfrom 120 to 180° C., a pressure of from 1 to 200 bar, preferably from 3to 150 bar and particularly preferably from 5 to 100 bar, and values ofthe liquid-side volumetric mass transfer coefficient k_(L)a of from 0.1s⁻¹ to 1 s⁻¹, preferably from 0. 2 s⁻¹ to 1 s⁻¹,

a) using a catalyst suspended in the reaction medium, where if a packedbubble column is employed this is operated in the upflow mode and at aratio of gas leaving the reaction vessel to gas fed to the reactionvessel of from 0.99:1 to 0.4:1, or

b) passing the liquid and gas in cocurrent in an upward directionthrough a fixed-bed reactor operated in the gas-circulation mode whilemaintaining a ratio of the gas fed to the reaction vessel to gas leavingthe reaction vessel of from 0.99:1 to 0.4:1.

The process of the present invention gives 1,4-butanediol in highspace-time yields together with high selectivity at a pressure below 200bar by means of a single-stage or multistage hydrogenation. In addition,long catalyst operating lives can be achieved.

The liquid-side volumetric mass transfer coefficient between the gasphase and the liquid phase k_(L)a is defined as

k_(L)a=k_(GL)×F_(GL),

where k_(GL) is the mass transfer coefficient for gas-liquid masstransfer and F_(GL) is the gas-liquid phase boundary area. The k_(L)avalue is, for example in Ullmanns Encyclopädie der technischen Chemie,Verlag Chemie, 4th edition (1973), Volume 3, pages 495 to 499, alsodescribed as the specific absorption rate.

The k_(L)a value is determined experimentally by measuring the hydrogenabsorption of a mixture of 50% by weight of butanediol and 50% by weightof water at the intended operating temperature. The procedure for theexperimental determination of k_(L)a has been described many times inthe specialist literature, for example in P. Wilkinson et al.: “MassTransfer and Bubble Size Distribution in a Bubble Column underPressure”, Chemical Engineering Science, Vol. 49 (1994) No. 9, pp.1417-1427, Ullmanns Encyclopadie der technischen Chemie, Verlag Chemie,Weinheim/Bergstr., 4th edition, 1973, Volume 3, pp. 495-499, H.Hoffmann: “Gepackte Aufstrom-Blasensäulen”, Chem.-Ing.-Tech. 54, (1982)No. 10, pp. 865-876 and A. Marquez et al.: “A Review of Recent ChemicalTechniques for the Determination of the Volumetric Mass-transferCoefficient k_(L)a in Gas-liquid Reactors”, Chemical Engineering andProcessing, 33 (1994) pp. 247-260.

According to the high k_(L)a values which are employed in carrying outthe process of the present invention, it is preferable to measure thehydrogen absorption under continuous operating conditions. As large aspossible a stream of the liquid mixture is fed in, hydrogen-free and ifappropriate together with suspended catalyst, at the desiredtemperature. The flow of the liquid mixture should be sufficiently highfor the liquid contents of the reactor to be replaced at least within 2minutes, preferably within 1 minute or less. At the same time,hydrogen-laden liquid mixture is taken off, depressurized to atmosphericpressure and the dissolved hydrogen thus liberated is determinedvolumetrically. The partial pressure of hydrogen in the gas phase islikewise measured.

The process of the present invention is preferably carried out usingtechnical-grade 1,4-butynediol which is in the form of an aqueoussolution and can additionally contain, as insoluble or dissolvedconstituents, components from the butynediol synthesis, eg. copper,bismuth, aluminum or silicon compounds. of course, it is also possibleto use butynediol which has been purified, eg. by distillation.Butynediol can be produced industrially from acetylene and aqueousformaldehyde and is customarily hydrogenated as a 30-60% strength byweight aqueous solution. However, hydrogenation can also be carried outin other solvents, for example alcohols such as methanol, ethanol,propanol, butanol or 1,4-butanediol. The hydrogen required for thehydrogenation is preferably used in pure form, but: it can also containfurther components such as methane and carbon monoxide.

According to the present invention, catalysts used are those which arecapable of hydrogenating C═C triple and double bonds to single bonds.They generally comprise one or more elements of transition groups I, VI,VII and VIII of the Periodic Table of the Elements, preferably theelements copper, chromium, molybdenum, manganese, rhenium, iron,ruthenium, cobalt, nickel, platinum and palladium. Particular preferenceis given to catalysts which comprise at least one element selected fromamong copper, chromium, molybdenum, iron, nickel, platinum andpalladium.

The metal content of these catalysts is generally 0.1-100% by weight,preferably 0.2-95% by weight, particularly preferably 0.5-95% by weight.

The catalyst preferably further comprises at least one element selectedfrom among the elements of main groups II, III, IV and VI, transitiongroups II, III, IV and V of the Periodic Table of the Elements and thelanthanides as promoter to increase the activity.

The promoter content of the catalyst is generally up to 5% by weight,preferably 0.001-5% by weight, particularly preferably 0.01-3% byweight.

As catalysts, it is possible to use precipitation, supported or Raneytype catalysts whose preparation is described, for example, in UllmannsEncyclopädie der technischen Chemie, 4th edition, 1977, Volume 13, pages558-665.

Support materials which can be used are aluminum oxides, titaniumoxides, zirconium dioxide, silicon dioxide, clays such asmontmorillonites, silicates such as magnesium or aluminum silicates,zeolites and activated carbons. Preferred support materials are aluminumoxides, titanium dioxides, silicon dioxide, zirconium dioxide andactivated carbons. Of course, mixtures of various support materials canalso serve as supports for catalysts which can be used in the process ofthe present invention.

These catalysts can be used either as shaped catalyst bodies, forexample as spheres, cylinders, rings and spirals, or in the form ofpowders.

Suitable Raney type catalysts are, for example, Raney nickel, Raneycopper, Raney cobalt, Raney nickel/molybdenum, Raney nickel/copper,Raney nickel/chromium, Raney nickel/chromium/iron or rhenium sponge.Raney nickel/molybdenum catalysts can be prepared, for example, by themethod described in U.S. Pat. No. 4,153,578. However, these catalystsare also sold by, for example, Degussa, 63403 Hanau, Germany. Forexample, a Raney nickel-chromium-iron catalyst is sold under the tradename Katalysator Typ 11 112 W® by Degussa.

When using precipitated or supported catalysts, these are reduced atfrom 150 to 500° C. in a stream of hydrogen or hydrogen/inert gas at thebeginning of the reaction. This reduction can be carried out directly inthe synthesis reactor. If the reduction is carried out in a separatereactor, the catalysts can be passivated on the surface at 30° C. usingoxygen-containing gas mixtures before being removed from the separatereactor. In this case, the passivated catalysts can be activated at 180°C. in a stream of nitrogen/hydrogen in the synthesis reactor beforebeing used, or can also be used without activation.

The catalysts can be used in a fixed bed or in suspension. If thecatalysts are in the form of a fixed bed, the reactor is, according tothe present invention, not operated in the customary downflow mode butusing an upward cocurrent of liquid and gas in such a way that theliquid and not the gas is present as the continuous phase.

Suspended catalysts are used in a particle size of generally 0.1-500 μm,preferably from 0.5 to 200 μm, particularly preferably from 1 to 100 μm.

If suspended catalysts are employed, then, when using packed bubblecolumns, the reaction is likewise carried out using an upward cocurrentof liquid and gas in such a way that the liquid and not the gas ispresent as the continuous phase. The ratio of gas leaving the reactionvessel to gas fed to the reaction vessel is, when using fixed-bedreactors and when using packed bubble columns with a catalyst suspendedin the reaction medium, from 0.99:1 to 0.4:1.

The ratio of gas leaving the reaction vessel to gas fed to the reactionvessel which is to be adhered to according to the present invention inthe case of fixed-bed reactors and in the case of catalysts suspended inthe reaction medium in packed bubble columns can be easily set by eithermetering in the appropriate amount of hydrogen as fresh gas or, aspreferred in industry, recirculating circulation gas and only making upthe loss of hydrogen resulting from chemical reaction and waste gas byfresh hydrogen.

The molar ratio of hydrogen to butynediol in the reactor is at least3:1, preferably from 4:1 to 100:1.

The process of the present invention is carried out over fixed-bedcatalysts in a gas-circulation mode, ie. the gas leaving the reactor iscirculated, if appropriate after being supplemented with fresh hydrogen,via a compressor back to the reactor. It is possible to convey the totalamount of circulation gas or a partial amount thereof via a jetcompressor. In this preferred embodiment, the circulation-gas compressoris replaced by an inexpensive nozzle. The work of compression isintroduced via the liquid which is likewise circulated. The increase inpressure of the liquid required to operate the jet compressor is fromabout 3 to 5 bar.

Suitable reactors for carrying out the process of the present inventionover fixed-bed catalysts are, for example, the fixed-bed reactor shownin FIG. 1 or a tube-bundle reactor as shown in FIG. 2.

FIG. 1 schematically shows the arrangement of a fixed-bed reactor whichcan be used in the process of the present invention. The reactor 1contains a bed of catalyst particles 2 having a mean diameter of fromabout 1 to 20 mm, preferably from 2 to 5 mm. To prevent the catalystparticles from being carried from the reactor, a wire mesh 3 is locatedat the upper end of the catalyst bed. The liquid feed 4 comprisingbutynediol and water is advantageously conveyed via the line togetherwith circulation liquid via line 5 as driving jet to a mixing nozzle 6in which fresh hydrogen via line 7 and circulation gas via line 8 aremixed in. A two-phase gas/liquid mixture 9 leaves the upper end of thereactor 1 and is separated in a gas/liquid separator 10. A substream 12of the gas stream 11 is taken off and discarded to avoid accumulation ofinert constituents. The circulation gas stream 8 is recirculated via acompressor 13 into the mixing nozzle 6. This compressor may be omittedif the circulation liquid 5 which is conveyed by the pump 21 can beprovided at sufficiently high pressure and the mixing nozzle 6 isdesigned as a jet compressor. A substream 14 of the circulation liquidis taken off as product stream. The heat of reaction liberated isremoved in the heat exchanger 16.

The process of the present invention can be carried out not only in theadiabatically operated fixed-bed reactor described in FIG. 1 but also inthe isothermally operated tube-bundle reactor described in FIG. 2.

FIG. 2 schematically shows the arrangement of a tube-bundle reactor inwhich the catalyst particles 2 having a mean diameter of from about 1 to20 mm, preferably from 2 to 5 mm, are arranged in the tubes 15.

The ratio of circulation liquid 5 to product 14 is, both in thefixed-bed reactor as shown in FIG. 1 and in the tube-bundle reactor asshown in FIG. 2, from 100:1 to 500:1, preferably 200:1. The diameter ofthe reactor is such that an empty-tube velocity of from 100 to 1000 m/his established for the liquid. The appropriate empty-tube velocity isdetermined for each type of catalyst in a laboratory apparatus. It isadvisable to set the empty-tube velocity at the maximum velocitypermissible with regard to catalyst abrasion. At empty-tube velocitiesabove about 1000 m/h, it has been found that for small catalystparticles there is an additional limitation set by the increasingpressure drop.

Main influencing parameters for fixing the empty-tube velocity are thecatalyst dimensions, its form and particle size distribution and itsabrasion behavior. A pressure drop figure of from about 0.02 to 0.15bar/m can, on the basis of experience, be used as a guide. The amount ofgas at the reactor outlet is preferably set such that the resultingempty-tube velocity is approximately comparable to the liquid empty-tubevelocity. However, it may be up to 90% lower.

For carrying out the process of the present invention using a catalystsuspended in the reaction medium, suitable reactors are jet nozzlereactors, stirred vessels and bubble columns with packing having apacking surface area of at least 500 m²/m³, preferably from 1000 to 2000m²/m³. Various types of jet nozzle reactors can be employed if they canensure, by means of a sufficiently high energy input which, on the basisof experience, is above 2 kW/m³, the high mass transfer from the gasphase to the liquid containing the suspended catalyst particles which isessential for the invention. Jet nozzle reactors which are equipped withan impulse exchange tube are particularly suitable. A widely distributedindustrial version of a jet nozzle reactor is, for example, the reactordescribed in EP-A 0 419 419. For energy input values of from 3 to 5kW/m³, this reactor still makes it possible to separate out the gasphase in simple separators without having to use additional equipmentsuch as foam centrifuges or cyclones.

FIG. 3 shows a jet nozzle reactor in which the liquid is conveyed vialine 5 via an external circuit having a heat exchanger 16 and drawshydrogen in in a driving jet compressor 6. To intensify the masstransfer, the two-phase mixture is conveyed via an impulse exchange tube17. At the lower end of the reactor 1, there may be a baffle plate 18which diverts the flow and makes the separation of the gas easier. Thegas rises in the outer annular space 19 toward the top and is againdrawn in by the driving jet compressor 6. The liquid from which the gashas been essentially separated is taken off at the lower end of thereactor, conveyed via a heat exchanger 16 to remove the heat of reactionand is again fed to the driving jet compressor 6.

Stirred vessels are suitable for carrying out the process of the presentinvention only when the energy input is in a range from 2 to 10 kW/m³.To convert the stirrer energy so as to achieve the high k_(L)a valuerequired by the invention, it is useful for the stirred vessel to haveinternal fittings which ensure the intimate mixing of gas and liquid,for example baffles.

In addition, the bubble columns provided with packing 20 shown in FIG. 4are also suitable for the process of the present invention. The surfacearea of the packing has to be at least 500 m²/m³, preferably from 1000to 2000 m²/m³. The packing 20 can be ordered or random, with orderedpacking as is known in terms of its geometry from distillativeseparation technology having the lowest pressure drop. Packingcomprising wire mesh, as is used in similar form in distillationtechnology, display particularly favorable properties. Examples are themesh packings Sulzer DX® or Sulzer Ex® which are sold by SulzerChemtechn., 8404 Winterthur, Switzerland.

The packing mentioned can also be coated directly with catalyticallyactive components. Such packing is described in EP-A 068 862, EP-A 201614 and EP-A 448 884. The fixed-bed reactor obtained using one of thesepacked bubble columns containing such packing has, for similarly highempty-tube velocities for the liquid and the gas of, in each case, from100 to 1000 m/h, preferably from 200 to 1000 m/h, the same high k_(L)avalues as in the suspension procedure.

The setting of the k_(L)a values according to the present invention offrom 0.1 s⁻¹ to 1 s⁻¹ which are decisive for simultaneously achieving ahigh selectivity and a high space-time yield, is carried out by means oftargeted technical measures tailored to the respective reactor type. Alltypes of reactor have in common an increased energy input compared withother ways of carrying out such a process. By means of specificstructural configurations and operating conditions, the energyintroduced is converted very effectively for the improvement of the masstransfer.

When the process of the present invention is carried out using suspendedcatalysts in stirred vessels, stirrer types having good gas-introducingproperties, for example disk stirrers or pitched blade stirrers as areknown, for example, from fermentation technology have to be used forsetting the k_(L)a values of from 0.1 s⁻¹ to 1 s⁻¹. The volumetricenergy input is from 2 to 10 kW/m³, with the lower value giving goodhydrogenation results only for small apparatuses. In the case of reactorsizes above about 0.5 m³, energy inputs of 5-10 kW/m³ are necessary. Instirred vessels, the energy is introduced via the drive power of thestirrer. These values for the energy input are higher than in the caseof customary gas-introduction reactions in stirred vessels, for examplefermentations or hydrogenations for which the energy input is from about0.2 to 1.0 kW/m³.

Jet nozzle reactors with suspended catalysts require volumetric energyinputs of more than 2 kW/m³, preferably 3-5 kW/m³. The energy isintroduced by means of the increase in the pressure of the liquid in thecirculation pump 21 in combination with the pressure reduction in thedriving jet compressor 6. Variation of amount circulated and pressurebuildup in the circulation pump enables the desired energy input to beset. The pressure buildup in the pump is usually in the range from 2 to5 bar.

If catalysts suspended in the reaction medium in packed bubble columnsare used in the process of the present invention, the surface area perunit volume of the packings has to be at least 500 m²/m³, but preferably1000-2000 m²/m³. To set the k_(L)a values of from 0.1 s⁻¹ to 1 s⁻¹ whichare required for the present invention, simultaneously high liquid andgas velocities of from 100 to 1000 m/h are required. In combination withthe indicated geometries of the packing, setting the indicated gas andliquid velocities ensures the necessary energy input which is however,as a result of the type of construction, lower than in the case ofstirred vessels or jet nozzle reactors. The achievement of theappropriate energy input which arises from the pressure drop of theflowing liquid and gas can be monitored by measuring the resultingpressure drop which is from 0.02 to 0.15 bar/m of packing. If necessary,the velocity of the liquid can be changed correspondingly in order toset the desired pressure decrease.

The values for the surface area per unit volume, the pressure decreasein the packing and the circulation rates for the liquid and gas givenabove for the suspension procedure in packed bubble columns also applyto packed bubble columns in which the packing itself is coated withcatalytically active material

When using fixed-bed reactors in the process of the present invention,the mean size of the catalyst particles has to be 1-20 mm, preferably2-5 mm, and the velocities of the liquid and gas flowing through have tobe 100-1000 m/h to set k_(L)a values of from 0.1 s⁻¹ to 1 s⁻¹. Thepressure drop established should be about 0.02-0.15 bar/m of fixed bed.

1,4-Butanediol is employed industrially in large amounts, eg. in thepreparation of THF or as a diol component in polyesters.

The process of the present invention is illustrated by means of thefollowing Examples. Unless otherwise indicated, use was made oftechnical-grade butynediol in the form of a 54% strength by weightaqueous solution which contained varying amounts of propynol. Theamounts of propynol correspond approximately to the amounts of propanolin the reaction product indicated in the Examples. The percentages inthe reaction products in the Examples are, unless otherwise indicated,percentages by weight calculated on an anhydrous basis which have beendetermined by gas chromatography.

EXAMPLES Example 1

A stirring autoclave having a liquid level of 130 ml and fitted with twobaffles, disk stirrers and built-in level control (a sintered metal fritfor holding back the catalyst) was charged with 10 g of Raney Ni/Mo (2%by weight of molybdenum, calculated as Mo, by impregnation of the Raneynickel with ammonium molybdate solution) in 50 ml of water and wassubsequently brought to 35 bar with hydrogen. By means of external oilheating, the internal reactor temperature was then brought to 140° C.and a hydrogen flow of 80 standard liters/h was set. The stirrer was setto 700 rpm, which ensured a k_(L)a of 0.2 s⁻¹. 100 g/h of a 54% strengthby weight aqueous butynediol solution were then pumped in. The internalreactor temperature rose to 149° C. The reaction product was obtained inan amount of 103 g/h and comprised 94.2% by weight of 1,4-butanediol,1.3% by weight of n-butanol, 3.3% by weight of n-propanol and a fewfurther products each in an amount of less than 0.08% by weight. The STYwas 0.4 kg of butanediol/l·h.

Example 2

Using a method similar to Example 1, 170 g/h of butynediol solution werehydrogenated over 10 g of Raney Ni/Mo (2.5% by weight of molybdenum,calculated as Mo). The initial temperature was 150° C. and the reactortemperature rose to 173° C. during the reaction. The product wasobtained in an amount of 176 g/h and comprised 92.4% by weight of1,4-butanediol, 0.4% by weight of 2-methylbutanediol, 2% by weight ofn-butanol and 4.7% by weight of n-propanol plus a few further productsin amounts of less than 0.08% by weight. The STY was 0.7 kg ofbutanediol/l·h.

Example 3

Using a method similar to Example 1, 60 g/h of butynediol solution werehydrogenated over 10 g of Raney Ni/Fe/Cr (type 11 112 W from Degussa).The liquid level in the reactor was 85 ml, the k_(L)a was 0.2 s⁻¹. Theinitial temperature was 140° C. and then rose to 144° C. The product wasobtained in an amount of 64 g/h and comprised 95.7% by weight ofbutanediol, 0.6% by weight of n-butanol and 1.8% by weight of n-propanolplus a few further products in amounts of less than 0.08% by weight. TheSTY was 0.25 kg of butanediol/l·h.

Example 4

Using a method similar to Example 1, 60 g/h of butynediol solution werehydrogenated at 600 rpm, corresponding to a k_(L)a of 0.1 s⁻¹, and areactor temperature of 105° C. The reactor output comprised 90% byweight of butanediol, 1.8% by weight of butenediol, 5% by weight of theacetal of 4-hydroxybutyraldehyde and butanediol, 2% by weight of4-hydroxybutyraldehyde and 4% by weight of butanol. After increasing thereactor temperature to 136° C., the conversion increased and thefollowing amounts of products were found in the output: 92% by weight ofbutanediol, 2.7% by weight of the acetal of 4-hydroxybutyraldehyde andbutanediol, 0.7% by weight of 4-hydroxybutyraldehyde and 3% by weight ofbutanol. Example 4 shows that the process of the present inventionenables higher selectivities to be achieved at higher conversions.

Example 5

A 400 ml oil-heated tube reactor having a diameter of 2.7 cm was filledwith 400 ml of 5 mm diameter Raschig rings made of metal mesh rings ofmaterial number 1.4541, steel list issued by Verein DeutscherEisenhuittenleute, 8th edition, Verlag Stahleisen mbH, Dusseldorf 1990,(UNS-No. S 32100). The tube reactor was installed in a reaction systemin which reaction liquid could be circulated via a gas/liquid separatorby means of a gear pump. The separator contained a filter through whichthe liquid and gas could be taken off continuously but which retainedthe catalyst. The feed of 200 g/h of butynediol as in Example 1 and 100standard l/h of fresh gas were fed in before the reactor. The reactorwas operated in the upflow mode. The space velocity of liquid was 170m³/m²h, the k_(L)a was 0.25 s⁻¹.

Before the reaction, the reaction system was charged in a similar way toExample 1 with 20 g of Raney Ni/Mo in 300 ml of water. At 30 bar and areactor temperature of from 145 to 151° C., the product was obtained inan amount of 213 g/h and comprised 93.3% by weight of butanediol, 0.3%by weight of 2-methylbutanediol, 1.5% by weight of n-butanol, 4.2% byweight of n-propanol and a few further products in amounts of less than0.08% by weight. The STY was about 0.25 kg of butanediol/l·h.

Example 6

The procedure of Example 5 was repeated with 5 g of Raney Ni/Mo beinginstalled. 100 g/h of butynediol solution were hydrogenated at a reactortemperature of about 122° C., 20 bar and 300 l/h of hydrogen. At a spacevelocity of liquid of 225 m³/m²·h and a k_(L)a of 0.3 s⁻¹ and completebutynediol conversion, n-butanol contents in the product of 2.2-2.7% byweight were obtained, with the 1,4-butanediol content being 77% byweight and the remainder being intermediates. After increasing the spacevelocity of liquid to 263 m³/m²h, corresponding to a k_(L)a of 0.4 s⁻¹,the n-butanol content fell to 1.3% by weight, the butanediol contentrose to 88% by weight.

Example 7

Using a method similar to Example 1, 60 g/h of 57% strength by weightaqueous butynediol solution (water content: 42% by weight) werehydrogenated at a reactor temperature of 127° C. and a k_(L)a of 0.2s⁻¹. After a reaction time of 125 hours, the product comprised 95.4% byweight of butanediol, 0.1% by weight of 2-methylbutanediol, 1.5% byweight of the acetal of 4-hydoxybutyraldehyde and butanediol, 2.6% byweight of butanol and 0.3% by weight of propanol. The reactortemperature was then increased to 141° C. Subsequently, the conversionof the intermediates also proceeded to completion and the selectivityrose. After an operating time of 173 hours, the following contents werefound: 98.8% by weight of butanediol, 0.1% by weight of2-methylbutanediol, 0.7% by weight of butanol and 0.3% by weight ofpropanol.

Comparative Example 1

Using a method similar to Example 1, distilled butynediol as a 50%strength by weight aqueous solution was hydrogenated. The pH of the feedsolution was adjusted to 6.6 by means of NaOH. At an oil bathtemperature of 140° C., an internal reactor temperature of 150° C. wasestablished at a feed rate of 100 g/h. After 24 hours of operation, thereactor output comprised 3% by weight of n-butanol, 0.5% by weight ofn-propanol and 96% by weight of 1,4-butanediol. After reducing thestirrer speed to 350 rpm, corresponding to a k_(L)a of 0.05 s⁻¹, thereactor temperature dropped to 141° C. and the reactor output comprised10% by weight of butynediol, 31% by weight of butenediol, 41% by weightof butanediol, 3% by weight of 4-hydroxybutyraldehyde, 0.5% by weight ofpropanol plus 4% by weight of butanol and 7% by weight of butenols. Theremainder was predominantly acetals.

Comparative Example 2

Using a method similar to Example 1, 100 g/h of technical-gradebutynediol solution were reacted at an oil bath temperature of 140° C.using 10 g of Raney Ni/Mo (1.8% by weight of molybdenum, calculated asMo). The internal reactor temperature was 149° C. The reaction producthad the following composition: 94.1% by weight of butanediol, 0.2% byweight of 2-methylbutanediol, 1.5% by weight of butanol and 4.2% byweight of propanol. After reducing the stirrer speed to 350 rpm,corresponding to a k_(L)a of 0.05 s⁻¹, the following hydrogenationresult was obtained: 40.3% by weight of butanediol, 37% by weight ofbutenediol, 2.1% by weight of butynediol, 3% by weight of butanol, 2.6%by weight of butenols, 3.1% by weight of 4-hydroxybutyraldehyde. Theremainder to 100% by weight comprised predominantly propanol, propenoland acetals of 4-hydroxybutyraldehyde with the diols.

We claim:
 1. A process for preparing 1,4-butanediol by continuouscatalytic hydrogenation of 1,4-butynediol in a stirred vessel, whichcomprises reacting 1,4-butynediol with hydrogen in the liquid continuousphase in the presence of a suspended hydrogenation catalyst at from 20to 300° C., a pressure of from 1 to 200 bar and values of theliquid-side volumetric mass transfer coefficient k_(L)a of from 0.1 s⁻¹to 1 s⁻¹.
 2. A process as claimed in claim 1 carried out at a pressureof from 3 to 150 bar.
 3. A process as claimed in claim 1 carried out ata pressure of from 5 to 100 bar.
 4. A process as claimed in claim 1,wherein the liquid-side volumetric mass transfer coefficient is from 0.2s⁻¹ to 1 s⁻¹.
 5. A process as claimed in claim 1, wherein the catalystused comprises at least one element selected from among the elements oftransition groups I, VI, VII and VIII of the Periodic Table of theElements.
 6. A process as claimed in claim 5, wherein the catalystcomprises at least one of the elements copper, chromium, molybdenum,manganese, rhenium, iron, ruthenium, cobalt, nickel, platinum andpalladium.
 7. A process as claimed in claim 5, wherein the catalystcomprises up to 5% by weight of at least one element selected from amongthe elements of main groups II, III, IV and VI, transition groups II,III, IV and V of the Periodic Table of the Elements and the lanthanides.8. A process as claimed in claim 5, wherein the catalyst comprises asupport selected from among the oxides of aluminum and titanium,zirconium dioxide, silicon dioxide, clays, silicates, zeolites andactivated carbon.